BACKGROUND OF THE INVENTION

The present invention relates to a catalytic process for aromatizing a C ₆ -C ₁₀ heavy naphtha feed to produce benzene, toluene and C ₈ and C ₉ aromatics, with the amount of xylene produced being maximized as compared to prior art methods.

Catalytic reforming is well known in the petroleum industry and refers to the treatment of naphtha fractions to improve the octane rating by the production of aromatics. One of the more important reactions occurring during reforming is the dehydrocyclization of acyclic hydrocarbons to aromatics. Catalytic reforming in this regard is also an important process for the chemical industry because of the great and expanding demand for aromatic hydrocarbons for use in the manufacture of various chemical products such as synthetic fibers, insecticides, adhesives, detergents, plastics, synthetic rubbers, pharmaceutical products, perfumes, drying oils, and various other products.

More specifically, for any given aromatic species, there exists fluctuating demands, notwithstanding that the total utilization of all aromatic hydrocarbons steadily increases. Aside from its use as a component of motor fuel, benzene serves as a starting material for the production of other aromatics such as styrene, phenol, synthetic detergents, DDT and nylon intermediates. Toluene is employed in aviation gasoline and as a high octane blending stock. As a petrochemical raw material, it is used in the production of solvents, gums, resins, rubber cement, vinyl organosols and other organic

chemicals. Mixed xylenes are primarily used in aviation gasoline and as a solvent for alkyd resins, lacquers, enamels and rubber cements, etc. More recently, however, para-xylene has been in great demand for use in the production of terephthalic acid employed in producing synthetic resins and fibers. To better meet this increased demand with respect to xylenes, and in particular para-xylenes, would be of great benefit to the industry.

Many different dehydrocyclization processes have been suggested in order to improve selectivity to aromatics and help meet the increased demand for aromatics. For example, in U.S. Patent No. 4,517,306, a dehydrocyclization process is disclosed which employs a type L zeolite, at least one Group VIII metal and an alkaline earth metal selected from the group consisting of barium, strontium and calcium. The catalyst is reduced in a hydrogen atmosphere at a temperature of from 480°C to 620°C prior to the reforming process. See also U.S. Patent Nos. 4,539,305 and 4,636,298.

U.S. Patent No. 4,650,565 discloses a dehydrocyclization process which involves contacting a naphtha feed in a reaction vessel with a dehydrocyclization catalyst comprising a large-pore zeolite containing at least one Group VIII metal to produce an aromatics product and a gaseous stream. The aromatics product is then separated from the gaseous stream and is passed through a molecular sieve which adsorbs paraffins present in the aromatic product, then the gaseous stream is used to strip the paraffins from the molecular sieve, and the gaseous stream and the paraffins are recycled to the reaction vessel. Preferably, the dehydrocyclization catalyst comprises a type L zeolite containing from 8%

to 15% by weight barium and from 0.6% to 1.0% by weight platinum, wherein at least 80% of the crystals of the type L zeolite are larger than 1000 Angstroms, and an inorganic binder selected from the group consisting of silica, alumina, and aluminosilicates.

U.S. Patent No. 4,158,025 discloses a process for selected aromatic hydrocarbon production. Selected aromatic hydrocarbon concentrates, such as benzene and mixed xylenes, are produced by way of a combination process which involves catalytic reforming followed by dealkylation. While the process affords some flexibility in the concentrate produced, it is particularly directed toward the maximization of benzene.

The dehydrocyclization or aromatization processes employed to date have generally focused on the production of benzene. A process which can improve the amount of xylenes produced, and in particular para- xylenes, by successfully and efficiently using heavier naphtha feeds would help to meet some of the increasing demand for such aromatics, and specifically para- xylenes, without being cost prohibitive.

Accordingly, it is an objective of the present invention to provide a process which efficiently improves the amount of xylene produced.

Another object of the present invention is to provide a process which is more selective toward the production of para-xylenes.

Yet another object of the present invention is to provide such a process which can efficiently use heavy naphtha feeds.

These and other objects of the present invention will become apparent upon a review of the following specifications and the claims appended thereto.

SUMMARY OF THE INVENTION

In accordance with the foregoing objectives, provided is a process for aromatizing a wide boiling range naphtha feed in order to produce a C ₈ stream with its C ₈ aromatics being rich in xylene. The process comprises aromatizing a wide boiling range naphtha feed, e.g., a C ₆ -C ₁₀ naphtha stream, over a high temperature treated L- zeolite catalyst in the potassium form containing a Group VIII metal. For the purposes of the present invention, a high temperature treated catalyst is defined as a catalyst that has been treated in an inert gas or reducing atmosphere at a temperature greater than or equal to 1025 ^Q F. Most preferably, the catalyst has been treated at a temperature in the range of from 1025°F to 1275°F while maintaining the water level of the effluent gas below 200 ppmv. Subsequent to the aromatization, a C ₈ stream rich in xylene is recovered.

Among other factors, the present invention is based on the surprising discovery that the use of a high temperature treated L-zeolite catalyst which is in the potassium form to aromatize a heavy naphtha feed produces a product stream that is significantly richer in xylene, and more specifically produces a C ₈ fraction which comprises largely xylenes, e.g., the C ₈ aromatics comprises at least 80% by weight xylene. The use of a

heavier naphtha feed in order to obtain a C ₈ stream so rich in xylenes is possible due to the particular catalyst employed in the process of the present invention. The catalyst used has been demonstrated to achieve a much higher selectivity to xylenes than has heretofore been deemed possible as compared to the use of L-zeolite aro atization catalysts presently in commercial use.

In a preferred embodiment, the catalyst used contains platinum in an amount ranging from about 1.0 to about 1.5 wt %. Such higher amounts of platinum further helps the activity and stability of the catalyst when processing a wide boiling range C ₆ -C _l0 naphtha feed.

DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENTS

In the process of the present invention, the catalyst used has been treated at a temperature in the range of from 1025°F to 1275°F while maintaining the water level of the effluent gas below 200 ppmv. This pretreatment has been found to permit the catalyst to operate with a heavier naphtha feed in order to produce a C ₈ stream rich in xylenes, i.e., wherein the C ₈ aromatics comprises at least 75% xylene, and most preferably at least 80% xylene.

The catalyst of the present invention is a large-pore zeolite containing at least one Group VIII metal, and is most preferably an L-zeolite. The preferred Group VIII metal is platinum, which is more selective for dehydrocyclization and which is more stable under reforming reaction conditions than other Group VIII metals. The catalyst should contain between 0.1% and 5% platinum based on the weight of the catalyst, more

preferably from 0.1% to 2.0%, and most preferably from about 1.0 to 1.5 wt %, e.g., about 1.2 wt %. The use of at least 1.0 wt % platinum is considered preferred in accordance with the present invention as it helps the activity and stability of the catalyst in working with the heavy naphtha feedstocks.

L-type zeolite catalysts are known. Examples of such L- zeolites are found, for example, in U.S. Patent No. 3,216,789, which is hereby incorporated by reference.

Type L zeolites are synthesized largely in the potassium form. These potassium cations, however, are exchangeable, so that other type L zeolites can be obtained by ion exchanging the type L zeolite in appropriate solutions. Generally, the zeolites have been exchanged in the prior art with an alkaline earth metal such as barium, strontium or calcium. The potassium has also been ion exchanged in the prior art with an alkali or alkaline earth metal, such as sodium, cesium or rubidium. For purposes of the present invention, however, it is important that the zeolite catalyst be substantially in the potassium form. For it has been discovered that the high temperature treated catalyst in the potassium form gives the desired increase in xylene production. Some exchange of ions can be allowed, but the amount of ions exchanged cannot be such as to materially effect the properties of the catalyst with respect to C ₈ xylene production.

An inorganic oxide can be used as a carrier to bind the large-pore zeolite. This carrier can be natural, synthetically produced, or a combination of the two. Preferred loadings of inorganic oxide are from 5% to 50% of the weight of the catalyst. Useful carriers include silica, alumina, aluminosilicates, and clays. The

preferred carrier is silica. The preparation of bound catalysts is described, for example, in U.S. Patent No. 4,830,732, which is hereby incorporated by reference in its entirety.

Preferably, the pretreatment process used on the catalyst occurs in the presence of a reducing gas such as hydrogen, as described in U.S. Patent No. 5,382,353, issued January 17, 1995, which is hereby expressly incorporated by reference in its entirety. Generally, the contacting occurs at a pressure of from 0 to 300 psig and a temperature of from 1025°F to 1275°F for from 1 hour to 120 hours, more preferably for at least 2 hours, and most preferably at least 4-48 hours. More preferably, the temperature is from 1050°F to 1250°F.

In general, the length of time for the pretreatment will be somewhat dependent upon the final treatment temperature, with the higher the final temperature the shorter the treatment time that is needed.

For a commercial size plant, it is necessary to limit the moisture content of the environment during the high temperature treatment in order to prevent significant catalyst deactivation. In the temperature range of from 1025°F to 1275°F, the presence of moisture is believed to have a severely detrimental effect on the catalyst activity, and it has therefore been found necessary to limit the moisture content of the environment to as little water as possible during said treatment period, to at least less than 200 ppmv.

In one embodiment, in order to limit exposure of the catalyst to water vapor at high temperatures, it is preferred that the catalyst be reduced initially at a temperature between 300°F and 700°F. After most of the

water generated during catalyst reduction has evolved from the catalyst, the temperature is raised slowly in ramping or stepwise fashion to a maximum temperature between 1025°F and 1250°F.

The temperature program and gas flow rates should be selected to limit water vapor levels in the reactor effluent to less than 200 ppmv and, preferably, less than 100 ppmv when the catalyst bed temperature exceeds 1025°F. The rate of temperature increase to the final activation temperature will typically average between 5 and 50°F per hour. Generally, the catalyst will be heated at a rate between 10 and 25°F/h. It is preferred that the gas flow through the catalyst bed during this process exceed 500 volumes per volume of catalyst per hour, where the gas volume is measured at standard conditions of one atmosphere and 60°F. In other words, greater than 500 gas hourly space volume (GHSV) . GHSV's in excess of 5000 h ^"1 will normally exceed the compressor capacity. GHSV's between 600 and 2000 h" ¹ are most preferred.

The pretreatment process occurs prior to contacting the reforming catalyst with a hydrocarbon feed. The large- pore zeolitic catalyst is generally treated in a reducing atmosphere in the temperature range of from 1025°F to 1275°F. Although other reducing gasses can be used, dry hydrogen is preferred as a reducing gas. The hydrogen is generally mixed with an inert gas such as nitrogen, with the amount of hydrogen in the mixture generally ranging from l%-99% by volume. More typically, however, the amount of hydrogen in the mixture ranges from about 10%-50% by volume.

The reducing gas entering the reactor should contain less than 100 ppmv water. It is preferred that it contain less than 10 ppmv water. In a commercial operation, the reactor effluent may be passed through a drier containing a desiccant or sorbent such as 4 A molecular sieves. The dried gas containing less than 100 ppmv water or, preferably, less than 10 ppmv water may then be recycled to the reactor.

In another embodiment, the catalyst can be pretreated using an inert gaseous environment in the temperature range of from 1025-1275°F, as described in copending

U.S. Serial No. (Attorney Docket No.

005950-442/T-5123) , filed May 25, 1995, which is hereby expressly incorporated by reference in its entirety.

The preferred inert gas used is nitrogen, for reasons of availability and cost. Other inert gases, however, can be used, such as helium, argon and krypton, or mixtures thereof.

The inert gas entering the reactor should contain less than 100 ppmv water. It is preferred that it contain less than 10 ppmv water. In a commercial operation, the reactor effluent may be passed through a drier containing a desiccant or sorbent such as 4 A molecular sieves. The dried gas containing less than 100 ppmv water or, preferably, less than 10 ppmv water may then be recycled to the reactor.

It is important, however, that the catalyst be reduced prior to the high temperature treatment in the inert atmosphere in the temperature range of from 1025 to 1275°F. For it has been found that simply heating the catalyst in a nitrogen atmosphere to elevated

temperatures can damage the catalyst. The catalyst must therefore, be first reduced at a temperature less than 1025°F. The reducing gas used is preferably hydrogen, although other reducing gases can also be used. The hydrogen is generally mixed with an inert gas such as nitrogen, with the amount of hydrogen in the mixture generally ranging from 1-99% by volume. More typically, however, the amount of hydrogen (or other reducing gas) in the mixture ranges from about 10-50% by volume.

In a preferred embodiment, the zeolite catalyst is reduced by contact with a reducing gas in a temperature range of from 300 to 900°F. Once the reduction of zeolite catalyst has taken place in the temperature range of from.300-900°F, the temperature can then be raised to the range of from 1025°F to 1275°F either stepwise or in a ramping fashion. The gaseous atmosphere is preferably inert in the temperature range from 900°F up to 1025°F. It is also preferred that the effluent gas water level be maintained below 200 ppmv in the temperature range of from 900°F to 1025°F.

In another preferred embodiment, the catalyst can be dried in an inert atmosphere such as a nitrogen atmosphere prior to reduction. The drying can take place while heating the catalyst from ambient

It is also necessary to limit the moisture content of the inert gaseous environment during the high temperature treatment in order to prevent significant catalyst deactivation. In the temperature range of from 1025°F to 1275°F, the presence of moisture is believed to have a severely detrimental effect on the catalyst activity, regardless of the type of environment in which the treatment takes place, and it has therefore been

found necessary to limit the moisture content of the environment to as little water as possible during said treatment period, to at least less than 200 ppmv.

In one embodiment, in order to limit exposure of the catalyst to water vapor at high temperatures, it is preferred that the catalyst be reduced initially at a temperature between 300°F and 700°F. After most of the water generated during catalyst reduction has evolved from the catalyst, the temperature is raised slowly in ramping or stepwise fashion to a maximum temperature between 1025°F and 1250°F. During the treatment in the temperature range of from 1025°F to 1250°F, the atmosphere is that of an inert gas.

The temperature program and gas flow rates when employing the inert gas atmosphere treatment should be selected to limit water vapor levels in the reactor effluent to less than 200 ppmv and, preferably, less than 100 ppmv when the catalyst bed temperature exceeds 1025°F. The rate of temperature increase to the final activation temperature can typically average between 5 and 50°F per hour. Generally, it is preferred that the catalyst be heated at a rate between 10 and 25°F/h. It is preferred that the gas flow through the catalyst bed (GHSV) during this process exceed 500 volumes per volume of catalyst per hour, where the gas volume is measured at standard conditions of one atmosphere and 60°F temperature.

The feed to the reforming process is typically a naphtha that contains at least some acyclic hydrocarbons or alkylcyclopentanes. This feed should be substantially free of sulfur, nitrogen, metals and other known poisons. These poisons can be removed by first using

conventional hydrofining techniques, then using sorbents to remove the remaining sulfur compounds and water. An example of a suitable feed, which feed has been hydrorefined to reduce sulfur content to acceptable levels, is shown below in Table 1:

TABLE 1 Feed Description

ASTM D 86 op

LV% St 145

10 184

30 198

50 219

70 243

90 262

EP 295 gravity, ° API 65.8

Carbon No. Distribution - wt% c ₅ 1.82 c ₆ 27.72 c ₇ 22.69 c ₈ 33.77

C ₉ 13.29

C _JO 0.72

PNA - Wt%

Paraffins (n+i) 72.32

Naphthenes 17.67

Aromatics 9.37

Unknown 0.64

Total 100.00

Because the catalyst of the present invention has been pretreated as previously described, it exhibits a longer run life with heavier feedstocks, e.g., containing at

least 5 wt % C ₉ + hydrocarbons, than similar catalysts having been subjected to a different treatment. The catalyst obtained via the treatment of the present invention, however, makes it quite practical to process feedstocks containing at least 5 wt % C ₉ + hydrocarbons, and for example at least 10 wt % C ₉ + hydrocarbons, with from 10-20 wt % C ₉ + hydrocarbons being preferred. Such larger amounts of C ₉ + hydrocarbons permits one to take better advantage of the present process as such a feedstock will have more C ₈ precursors to permit even greater amounts of xylene to be produced.

The feed can be contacted with the catalyst in either a fixed bed system, a moving bed system, a fluidized system, or a batch system. Either a fixed bed system or a moving bed system is preferred. In a fixed bed system, the preheated feed is passed into at least one reactor that contains a fixed bed of the catalyst. The flow of the feed can be either upward, downward, or radial. The pressure is from about 1 atmosphere to about 500 psig, with the preferred pressure being from abut 50 psig to about 200 psig. The preferred temperature is from about 800°F to about 1025°F. The liquid hourly space velocity (LHSV) is from about 0.1 hr ^{" 1} to about 10 hrs" ¹ , with a preferred LHSV of from about 0.3 hr" ¹ to about 5 hrs' ¹ . Enough hydrogen is used to insure a H ₂ /HC ratio of up to about 20:1. The preferred H ₂ /HC ratio is from about 1:1 to about 7:1, and most preferred 2:1 to about 6:1. Reforming produces hydrogen. Thus, additional hydrogen is not needed except when the catalyst is reduced and when the feed is first introduced. Once reforming is underway, part of the hydrogen that is produced is recycled over the catalyst.

Once the reforming process has been completed, the product can be separated into the desired streams or fractions. A C ₈ aromatics stream is recovered using conventional techniques such as distillation and/or extraction.

In a preferred embodiment, for example, the effluent from the aromatization reactor is cooled and then sent to a separator where the liquid is recovered for further processing. The gas from the separator is collected and part is recycled to the reactor inlet and part is excessed from the aromatization process. The liquid which contains mostly C ₈ -C ₉ aromatics and some C ₁₀ aromatics as well as non-aromatics is depentanized and then sent to a first distillation column to recover benzene and toluene as an overhead product. This overhead cut which also includes non-aromatics in the benzene and toluene boiling range is further processed in an aromatics extraction plant, e.g., Udex, Sulfolane, Krupp, to separate the aromatics from the non-aromatics. The aromatics stream is then sent to a second distillation column where the benzene is recovered as an overhead product and the toluene as a bottoms product.

The bottoms stream from the first distillation column, which consists primarily of C ₈ and C ₉ aromatics, and some C _j o aromatics is sent to a third distillation column where the C _g aromatics are removed as an overhead product and the C ₉₊ aromatics as a bottoms product. The C ₈ aromatics stream from the third distillation column is processed in a plant where the PX (para-xylene fraction) is removed by adsorption, or crystallization, or a combination of adsorption-crystallization. The PX lean stream from this process, also known as the raffinate stream, is sent to a xylene isomerization plant where

the xylenes are reacted to equilibrium thereby converting some of the meta and ortho-xylenes to para- xylenes. The resulting effluent from the xylene isomerization plant is distilled to remove light (benzene and toluene) and heavy aromatics (C ₉₊ ) . The remaining xylenes are then recycled to the third distillation column for further processing in the PX removal plant.

It has been found that the process of the present invention results in at least about a 27% increase in xylene production. Moreover, the present process is more selective in that the ethylbenzene yield is generally reduced by about 40% which results in a C ₈ aromatics distribution which contains at least 80% xylenes.

The invention will be illustrated in greater detail by the following specific examples. It is understood that these examples are given by way of illustration and are not meant to limit the disclosure or the claims to follow. All percentages in the examples, and elsewhere in the specification, are by weight unless otherwise specified.

EXAMPLE 1 Preparation of 0.6% Pt/K-Ba L Zeolite Catalyst:

Silica bound zeolite extrudates were ion-exchanged with barium. The barium exchanged extrudates were then dried and calcined at a temperature exceeding 1100°F. The calcined extrudates were pore-fill impregnated using platinum tetra-ammonium di-chloride solution. The resulting extrudate was dried and calcined at 500-550°F in a steam/air environment and then crushed and sieved

to a particle size of 14-28 mesh. The resulting catalyst was 0.6 wt.% Pt/K-Ba L zeolite.

EXAMPLE 2

Preparation of 1.2% Pt/K-Ba L Zeolite Catalyst:

A 1.2% Pt/K-Ba L zeolite silica-bound extrudate was prepared in a similar manner to that described in Example 1. Thus, following the barium ion-exchange and calcination as in Example 1, the catalyst was impregnated with platinum as described in Example 1, except that sufficient platinum tetra-ammoniu di- chloride was used to achieve a platinum content of 1.2 wt.%. Furthermore, following the platinum impregnation step, the catalyst of the present example was not calcined but dried in air at 185°F. Prior to use in the reactor, the catalyst was crushed and sieved to a particle size of 14-28 mesh.

EXAMPLE 3

Preparation of 1.2% Pt/K L Zeolite Catalyst

A silica-bound K L zeolite extrudate was impregnated with platinum without a barium ion exchange step. The platinum impregnation step was as described in Example 2 to obtain a 1.2% Pt/K L zeolite. The resulting extrudate was then dried in air at 185°F. Prior to use in the reactor, the catalyst was crushed and sieved to a particle size of 14-28 mesh.

EXAMPLE 4

Runs were carried out involving the catalysts of Examples 1-3. All the runs were carried out as follows:

Four cubic centimeters of 14-28 mesh silica-bound catalyst were charged to a 1/2" diameter reactor. The catalyst was then dried in an inert gas atmosphere, reduced in the presence of hydrogen and then subjected to a high temperature treatment in a reducing atmosphere as follows:

1. Nitrogen flow was established over the catalyst at room temperature and atmospheric pressure. The nitrogen flow was such as to obtain a GHSV of 1000.

2. The unit pressure was brought up to 50 psig and set to control at that pressure.

3. The catalyst was next heated from ambient temperature to 300°F at a rate of 50°F/hr.

4. On reaching 300°F the reactor temperature was held at 300°F for two hours.

5. The catalyst was then heated to 500°F at 25°F/hr.

6. On reaching 500 F, this condition was held for two hours while flowing nitrogen over the catalyst.

7. After a two hour hold the nitrogen flow was switched off and hydrogen flow was started at 1000 GHSV. Hydrogen was fed on a once-through basis at 500°F for one hour. 8. One hour after introducing hydrogen to the reactor, the catalyst temperature was increased to 900°F at a rate of 10°F/hr.

9. For the high temperature treatment, the temperature was increased 10°F/hr until the target treatment

temperature was reached, then this final temperature was held for 3 hours.

10. After three hours at the treatment temperature, the temperature of the catalyst/reactor was reduced to the desired starting temperature and the reactor pressure was adjusted to the desired operating pressure.

All of the catalysts were tested at the following operating conditions: Reactor pressure - psig 75 LHSV-hr-1 1

H2/HC mol/mol 5

Severity 70% Aromatics as wt% of feed.

After completing the high temperature treatment, the reactor temperature was reduced to about 800°F and the pressure was established at 75 psig. The naphtha feed, which was the same as described in Table 1, was then introduced at 4 cc/hr. After equilibrating, reactor inline samples were taken and analyzed by a gas chro atograph to determine the aromatics concentration in the reactor effluent. The reactor temperature was subsequently increased until the aromatics concentration of the reactor effluent was 70%. After achieving this target temperature, the reactor temperature was increased in small increments as necessary to maintain the target severity of 70% aromatics in the reactor effluent. At the same time the ethylbenzene and xylene concentration was determined as well as the C ₈ aromatics and xylene distribution.

The results of the runs are shown in Table 2 below:

TABLE 2

0.6 t% 1.2 t% 1.2 t% Pt /K-Ba Pt /K-Ba Pt /K

Cg Aromatics ethyl- Distribution benzene 33.6 31.9 17.9 xylene 66.4 68. 1 82 . 1

Xylene distribution para-xylene 12.3 10.3 14.4 meta-xylene 39.8 33.8 44.3 ortho-xylene 47.9 55 .9 41.3

Yields of Aromatic Products -

Benzene 23.9 24. 3 22 . 6 Toluene 26.4 28.0 29. 0

Xylene 11.6 10.8 14.0

Ethyl-benzene 5.9 5.0 3. 1

C, 4.6 3 .3 2 .4 The foregoing results demonstrate that the process of the present invention (using the catalyst of Example 3) , resulted in a significant increase in xylene production (and concurrently a reduction in ethylbenzene production) and an important improvement in para-xylene distribution. This result is quite surprising and can be important in meeting the increased demand of chemicals such as xylene, and in particular, para- xylenes.

While the invention has been described with preferred embodiments, it is to be understood that variations and modifications may be resorted to as will be apparent to those skilled in the art. Such variations and modifications are to be considered within the purview and the scope of the claims appended hereto.